Effective de-scaling for desalination plants and a new brine-forward multi-stage flash concept

ABSTRACT

This invention presents an innovative de-scaling step to effectively eliminate scale obstacles in desalination plants to: (1) allow better performance ratio by either reducing steam load or increasing distillate production, and thus more efficient and cost effective plants; (2) reduce the enormous volumes of the total required seawater (recycle brine and returned cooling seawater) and reject brine; (3) eliminate the addition of scale inhibitors and frequent shut-downs to remove scale buildup; and (4) minimize reduction in heat transfer coefficients. This invention, as a result of the effective de-scaling step, also presents a new effective design for Multi-Stage Flash (MSF) desalination plants, which is based on what its coined a Brine-Forward (BF) concept. The BF-MSF concept is multi-boiling without supplying additional heat after the train&#39;s brine heater, and multi-concentration without supplying additional seawater after the first train of the plant. Thus, the BF-MSF concept entirely eliminates brine recycling and returned (rejected) cooling seawater, which provides tremendous savings in pumping power.

RELATED APPLICATION

This application is related to U.S. patent application Ser. No.13/066,841, filed Apr. 26, 2011, now U.S. Pat. No. 8,915,301.

BACKGROUND OF TELE INVENTION

For example, two seawater desalination plants coupled with two powerplants are recently installed in the Arabian Gulf area. The desalinationplants are nearly identical and based on the Recycle-Brine Multi-StageFlash (RB-MSF) concept. Each plant consists of eight identical trains,and each train produces 15 million U.S. gallons per day (MGD) ofdistillate at 100% load. The distillate production capacity of eachplant is thus 120 MGD. FIG. 1 shows a simplified flow diagram for eachBR-MSF train. The train mainly consists of three major sections: (1)heat rejection section; (2) heat recovery section; and (3) brine heater.The number of stages in the heat rejection section is 3 while the numberof stages in the heat recovery section is 20.

Seawater is pretreated by mainly screening and chlorination beforeentering the last stage of the heat rejection section through heattransfer tubing (condensers/pre-heaters). A significant volume ofpretreated seawater is required for cooling in the heat rejectionsection. About 70% of the preheated seawater exiting the heat rejectionsection is rejected as “returned cooling seawater”. About 20% of suchreturned cooling seawater is recycled for mixing with seawater feedstream to maintain a constant temperature at the entrance of the heatrejection section whereas the remaining returned cooling seawater ismixed with reject brine from the last stage of the heat rejectionsection before blowing it back to the sea.

The remaining portion of the preheated seawater existing the heatrejection section (about 30% of the total required seawater) is mixedwith additives (e.g., anti-scale and anti-foam) and fed to a vacuumde-aerator. The preheated and de-aerated seawater is then mixed with aportion of the flashed off brine from the last stage of the heatrejection section to form recycle brine. The recycle brine is mixed withan oxygen scavenger before feeding it to the last stage of the heatrecovery section.

In the heat recovery section, recycle brine passes through heat transfertubing (condensers/pre-heaters) before entering the brine heater. Thebrine heater is externally driven by low-pressure (LP) andintermediate-pressure (IP) steam to heat recycle brine to the desired ordesigned top brine temperature of the train.

The heated recycle brine flows from the brine heater to the firstflashing stage where the pressure is lowered so that it is just belowthe vapor saturation pressure of water. This sudden introduction ofheated recycle brine into a lower pressure stage causes it to boil sorapidly as to flash into vapor. However, a small fraction of the recyclebrine is converted into vapor. The remaining recycle brine thus passesthrough a series of flashing stages, each possessing a lower pressure tolower the boiling point of recycle brine than the previous stage. Thisallows consecutive reduction of the boiling point of recycle brine as itgets more concentrated in going down the flashing stages. MSF is thus amultiple boiling concept without supplying additional heat after thebrine heater.

The flashed off vapor condenses on the tubes side of thecondensers/pre-heaters and transports across the heat recovery sectionas distillate. The released latent heat of the condensed vapor is usedto preheat the incoming recycle brine in the heat recovery section.Since the colder recycle brine entering the heat recovery sectioncounter flows with the flashed off brine/distillate, relatively littleheat energy leaves in the flashed off brine from the last stage of theheat recovery section. The unflashed portion of recycle brine thenpasses through additional flashing stages in the heat rejection sectionto recover more distillate. As such, most of the heat energy of theflashing process is exchanged with the recycle brine (heat recoverysection) and seawater (heat rejection section) flowing into the train.

However, thermally unstable ions (bicarbonate, magnesium and calcium) inseawater present engineering difficulties that severely limit thepotential capacity of any desalting method. In the RB-MSF train,bicarbonate in seawater is thermally and partially broken down inpre-heaters (heat rejection section) into carbon dioxide and hydroxides.Carbon dioxide is removed from seawater by the vacuum de-aerator whereasthe released hydroxides along with the gradual temperature rise ofseawater (heat rejection section) and recycle brine (heat recoverysection) could trigger magnesium hydroxide (brucite) scaling.

Calcium, on the other hand, has two possible forms of sulfate (anhydrousand/or hemihydrate) at the top brine temperatures (110° C.). Thesolubility limits of calcium sulfate anhydrous or hemihydrate areinversely and steeply proportional with temperatures. Scale inhibitorsare added to delay the precipitation of sulfate scales even though theireffect is very limited in solving such scales especially within thefront end flashing stages in the heat recovery section. The temperaturetolerance limit of scale inhibitors also dictates the train's top brinetemperature (e.g., 90° C. for polyphosphates and 110° C. forpolycarboxylates or polymeric).

The desalination plants (I and II) are unusually located near a marshyshallow seawater area. Seawater can not be drawn from a deep watercolumn to take advantage of reduced oxygen, suspended solids andmicrobial activity at depth. Seawater in that area is also known ofhaving high silt index and low natural current. As shown in Table 1, theconsequences of the plants' location along with discharging back to thesea a copious volume of reject brine (905.4 MGD) from both plants areclearly pronounced in the differences of the TDS and sulfate levelsbefore and after operating the plants. Sulfate level, for example, hassurged from 2,700 mg/L in 2006 to 4,100 mg/L in 2013 (increased by 52%).

Table 2 shows different operating conditions for a given RB-MSF train tocope with the consequences of improper plants' location. FIG. 2 revealsthat the concentration of calcium sulfate anhydrous in the front-endflashing stages close to the brine heater (stage 1 to 4) when seawaterhas a normal sulfate level (2,700 mg/L) in 2006 is slightly below thesaturation line. If the same operating conditions of the RB-MSF train asbuilt including the same concentration factors (recycle brine/seawaterand reject brine/seawater) are held regardless of the changes inseawater composition (4,100 mg/L of sulfate level in 2013), theconcentration of calcium sulfate anhydrous in the flashing stages closeto the brine heater is over saturated (FIG. 2). As presented in Table 2,the TDS and sulfate levels are appreciably increased in recycle brineand reject brine (e.g., the TDS and sulfate surges in recycle brine are,respectively, from 56,101 to 70,943 mg/L and from 3,835 to 5,823 mg/L).

In order to operate the RB-MSF plants properly and manage thedetrimental changes in seawater composition (Table 1) in the form ofsulfate scales, the reject brine concentration (C_(B)) must be fixed at63,200 mg/L of TDS as the plants originally designed. The concentrationfactor of recycle brine/seawater (C_(r)/C_(F)) would be reduced to 1.13from 1.42 whereas the concentration factor of reject brine/seawater(C_(B)/C_(F)) would be reduced to 1.27 from 1.6 (Table 2). Theconcentration of calcium sulfate anhydrous in the flashing stages closeto the brine heater would then be kept at the saturation border line asshown in FIG. 2. However, the required seawater for blending withflashed off brine to form recycle brine would increase by 75% (Table 2).This, in turn, would: (1) elevate the required amounts of additives(anti-foam, anti-scale, and oxygen scavenger); (2) require larger vacuumde-aerators; (3) substantially increase the volume of reject brine (by121%); and (4) require higher pumping power. It should be pointed outthat the total required seawater (returned cooling seawater andpreheated seawater to be blend with flashed off brine from the heatrejection) remains the same as in the original RB-MSF design (Table 2);returned cooling seawater is just proportionally decreased with theincrease of preheated seawater to feed the heat recovery section.

Recycle Brine (RB) became the conventional approach compared to theOnce-Through (OT) approach in designing MSF plants in the past 20 yearsfor several presumed reasons. First, recycle brine reduces the actualvolume of seawater to feed the trains for distillate production, whichwould reduce the amounts of additives, size of vacuum de-aerators andvolume of reject brine. Second, RB-MSF adds a heat rejection section ineach train to: (1) recover more distillate at the low flashing range(40-33° C.); (2) maintain a constant temperature at the entrance of theheat rejection section by blending a portion of returned coolingseawater with seawater feed stream, (3) dilute reject brine with theremaining portion of the returned cooling seawater before blowing itdown to the sea; and (4) eliminate the conventional seawaterde-alkalization step. The conventional seawater pretreatment indesalination plants typically includes a de-alkalization step by dosingan acid (e.g., sulfuric acid) to convert bicarbonate to carbon dioxide,removing carbon dioxide by the de-aerator, and then neutralizing thepretreated seawater or distillate (e.g., with caustic soda) to re-adjustthe pH.

On the other hand, the OT-MSF is a simpler approach than the RB-MSFapproach since it consists of a number of stages and a brine heater asshown in FIG. 3. The performance ratio (PR) of OT-MSF train iscomparable to RB-MSF train for the same number of stages and top brinetemperature. Reject brine from OT-MSF train has low TDS gains comparedto seawater and is directly discharged to the sea. As such, OT-MSF traindoes not require a substantial volume of cooling seawater since it hasno heat rejection section nor a substantial pumping power to circulate atremendous volume of recycle brine as is the case with RB-MSF plant.However, OT-MSF plant has a lower ratio of distillate/seawater thanRB-MSF plant (no recycle brine), and it is brine from the last stage istypically rejected at 40° C. rather than 33° C. (a further but slightloss of distillate recovery), which are considered the drawbacks of theOT-MSF approach.

When an OT-MSF train, however, is applied to the changed composition ofseawater (Table 1) by fixing C_(B) at 63,200 mg/L of TDS as the RB-MSFtrain originally designed, it provided better performance than RB-MSFtrain with distinct advantages (Table 2). Such advantages can be seenin: (1) reducing the steam load (from 62.9 to 33.0 kg/s) and thereforesubstantially increasing the PR (from 9.4 to 13.6); (2) entirelyeliminating the substantial volume of cooling seawater; and (3) keepingthe concentration of calcium sulfate anhydrous in the flashing stagesclose to the brine heater slightly below the saturation border line(FIG. 2).

Regardless of the type of MSF plants, whether they are based on RB or OTapproach, in order to produce 240 MGD of distillate and operate theplants properly, the saturation envelop of calcium sulfate anhydrousmust be controlled by fixing C_(B) at 63,200 mg/L of TDS. Asconsequences, the volumes of the total required seawater (2,247 MGD forRB-MSF and 1,146 MGD for OT-MSF) and reject brine (905.4 MGD for bothRB-MSF and OT-MSF) are enormous. Thermal energy to brine heaters can bereduced from one-half to two thirds by pairing with waste heat from theturbines of the coupled power plants and at the same time providingcooling for the power plants, but pumping energy to circulate enormousvolumes of seawater, recycle brine and reject brine along withpre-treating a large volume of seawater have become the largestoperating costs of MSF plants. The discharge of a copious volume ofreject brine to the sea increases TDS around seawater intake lines sinceit is dispersion not fast enough (shallow seawater, low natural current,and insufficient or absent mechanical dispersion devices), whichdeteriorates the natural composition of seawater and imposes differentsets of plants' operating conditions. Disposal of reject brine has alsoa significant environmental impact on marine habitat including depletedoxygen, residue of the deoxygenating agent, concentrated toxic species(e.g., derivatives of boron and chlorine), and induction of gypsumprecipitation. Such enormous engineering, economic and environmentaldifficulties are solely caused by calcium sulfate scale; a seriousobstacle to the development of low cost distillate production fromseawater desalination plants.

Effective de-scaling is thus essential in seawater desalination since itwould: (1) allow better performance ratio (ratio of distillate toheating steam) and thus more efficient and cost effective plants; (2)reduce the volumes of seawater (substantial savings on pumping power andpre-treatment) and reject brine; (3) eliminate the addition of scaleinhibitors; (4) allow desalination plants such as MSF to reach a higherrange of top brine temperatures (120-150° C.) and thus increase theirperformance ratio by either decreasing steam load or increasingdistillate production; (5) prevent maintenance shut-downs for chemicaland mechanical scale removal; and (6) slow the decrease of heat transfercoefficients.

Thus, the first objective of this invention is to effectively eliminatescale issues in any desalination plants to achieve the above mentionedbenefits. As a result of effective de-scaling, the second objective ofthis invention is to develop a new effective design for MSF plants,which is based on what I coined Brine-Forward (BF), to entirelyeliminate brine recycling and returned cooling seawater.

SUMMARY OF THE INVENTION

In one aspect, the present invention provides a method for separatingbrucite and sulfate from a sulfate-rich saline stream to producede-scaled saline stream suitable for a desalination system. Theinventive method comprises the steps of: (a) separating brucite from thesulfate-rich saline stream to produce de-brucited saline stream; (b)separating sulfate from the de-brucited saline stream to producede-scaled saline stream; (c) de-aerating the de-scaled saline stream bya vacuum de-aerator or an indirect-contact heat exchanging stripper toproduce de-aerated saline stream; and (d) feeding the de-aerated salineto the desalination system to produce distillate and reject brine.Brucite is separated from the sulfate-rich saline stream in step (a) by:(i) mixing the sulfate-rich saline stream with an appropriate amount ofa hydroxide source of an inorganic-based additive or an organic-basedadditive to form precipitates comprising brucite in a first precipitatorunit; and (ii) removing precipitates from the sulfate-rich saline streamby an appropriate filtration unit to produce de-brucited saline streamand dewatered precipitates. Sulfate is separated from the de-brucitedsaline stream in step (b) by: (i) mixing the de-brucited saline streamwith appropriate amounts of an aluminum-based or iron-based additive,and a hydroxide source of an inorganic-based additive or anorganic-based additive to form precipitates comprising calciumsulfoaluminate or calcium sulfoferrate in a second precipitator unit;and (ii) removing precipitates from the de-brucited saline stream by anappropriate filtration unit to produce the de-scaled saline stream anddewatered precipitates. Sulfate-rich saline stream is seawater, brinefrom seawater desalination plants, natural brine, brackish water,produced water, flue gas desulphurization spent water, agriculturaldrainage water, acid mine drainage water, mineral slurry transportwater, paper mills spent water, aluminum anodizing spent water, spentwater from fertilizer production, lime slaking, or spent water fromtextile production. The desalination system can be a Recycle-BrineMulti-Stage Flash Desalination, Once-Through Multi-Stage FlashDesalination, Multi-Effect Distillation, Thermal Vapor Compression,Mechanical Vapor Compression, Vacuum Membrane Distillation,Direct-Contact Membrane Distillation, Osmotic Membrane Distillation,Reverse Osmosis, Forward Osmosis, or a combination of such methods. Thehydroxide source of the inorganic-based additive is calciumchloroaluminate, calcium chloroferrate, lime, and hydrated lime. Thehydroxide source of the organic-based additive is ammonia, methylamine,ethylamine, isopropylamine, propylamine, dimethylamine, diethylamine,diisopropylamine, and dipropylamine. The aluminum-based additive iscalcium chloroaluminate, aluminum chlorohydrate, calcium aluminate,sodium aluminate, aluminum acetate, aluminum chloride, and aluminumnitrate. The iron-based additive is calcium chloroferrate, calciumferrate, sodium ferrate, iron chloride, and iron nitrate.

In yet another aspect, the present invention provides a method forpartially preheating, and separating brucite and sulfate from asulfate-rich saline stream to produce de-scaled saline stream suitablefor the Brine-Forward Multi-Stage Flash (BF-MSF) desalination system.The inventive method comprises the steps of: (a) partially preheatingsulfate-rich saline stream; (b) separating brucite from the partiallypreheated sulfate-rich saline stream to produce de-brucited salinestream; (c) separating sulfate from the de-brucited saline stream toproduce de-scaled saline stream; (d) de-aerating the de-scaled salinestream by a vacuum de-aerator or an indirect-contact heat exchangingstripper to produce de-aerated saline stream; (e) feeding the de-aeratedsaline stream to the first train of the BF-MSF desalination system toproduce a first distillate and a first brine; (f) applying a brineforward feeding procedure to feed the first brine from the first trainof the BF-MSF desalination system to a second train to produce a seconddistillate and a second brine; and (g) repeating the brine forwardfeeding procedure in subsequent trains of the BF-MSF desalination systemto allow consecutive increments in brine’ total dissolved solids untilthe total dissolved solids reach a desired level in the last train ofthe BF-MSF desalination system to reject brine. The sulfate-rich salinestream is partially preheating in step (a) by: (i) passing a portion ofthe sulfate-rich saline stream through heat transfer tubings of the heatrejection section in the last train of the BF-MSF desalination system toexchange heat energy with flashing brine in the same heat rejectionsection to produce preheated portion of sulfate-rich saline stream; and(ii) blending the preheated portion of sulfate-rich saline stream withanother portion of sulfate-rich saline stream to produce partiallypreheated sulfate-rich saline stream. Brucite is separated from thepreheated sulfate-rich saline stream in step (b) by: (i) mixing thepreheated sulfate-rich saline stream with an appropriate amount of ahydroxide source of an inorganic-based additive or an organic-basedadditive to form precipitates comprising brucite in a first precipitatorunit; and (ii) removing precipitates from the preheated sulfate-richsaline stream by an appropriate filtration unit to produce de-brucitedsaline stream and dewatered precipitates. Sulfate is separated from thede-brucited saline stream in step (c) by: (i) mixing the de-brucitedsaline stream with appropriate amounts of an aluminum-based oriron-based additive, and a hydroxide source of an inorganic-basedadditive or an organic-based additive to form precipitates comprisingcalcium sulfoaluminate or calcium sulfoferrate in a second precipitatorunit; and (ii) removing precipitates from the de-brucited saline streamby an appropriate filtration unit to produce the de-scaled saline streamand dewatered precipitates.

In yet another aspect, the present invention provides a method forpreheating, and separating brucite and sulfate from a sulfate-richsaline stream to produce de-scaled saline stream suitable for theBrine-Forward Multi-Stage Flash (BF-MSF) desalination system. Theinventive method comprises the steps of: (a) preheating the sulfate-richsaline stream; (b) separating brucite from the preheated sulfate-richsaline stream to produce de-brucited saline stream; (c) separatingsulfate from the de-brucited saline stream to produce de-scaled salinestream; (d) de-aerating the de-scaled saline stream by a vacuumde-aerator or an indirect-contact heat exchanging stripper to producede-aerated saline stream; (e) feeding the de-aerated saline stream tothe first train of the BF-MSF desalination system to produce a firstdistillate and a first brine; (f) applying a brine forward feedingprocedure to feed the first brine from the first train of the BF-MSFdesalination system to a second train to produce a second distillate anda second brine; and (g) repeating the brine forward feeding procedure insubsequent trains of the BF-MSF desalination system to allow consecutiveincrements in brine’ total dissolved solids until the total dissolvedsolids reach a desired level in the last train of the BF-MSFdesalination system to reject brine. Sulfate-rich saline stream ispreheated in step (a) by: (i) passing a portion of sulfate-rich salinestream through heat transfer tubings of the heat rejection section ofthe last train of the BF-MSF desalination system to exchange heat energywith flashing brine in the same heat rejection section to produce afirst preheated portion of sulfate-rich saline stream; (ii) preheatinganother portion of sulfate-rich saline stream by a direct-contact heatexchanger using a source of waste heat energy or a flashing fluid or aheating fluid or a fluid requires condensing to produce a secondpreheated portion of sulfate-rich saline stream; and (iii) blending thefirst preheated portion of sulfate-rich saline stream with the secondpreheated portion of sulfate-rich saline stream to produce preheatedsulfate-rich saline stream. Brucite is separated from the preheatedsulfate-rich saline stream in step (b) by: (i) mixing the preheatedsulfate-rich saline stream with an appropriate amount of a hydroxidesource of an inorganic-based additive or an organic-based additive toform precipitates comprising brucite in a first precipitator unit; and(ii) removing precipitates from the preheated sulfate-rich saline streamby an appropriate filtration unit to produce de-brucited saline streamand dewatered precipitates. Sulfate is separated from the de-brucitedsaline stream in step (c) by: (i) mixing the de-brucited saline streamwith appropriate amounts of an aluminum-based or iron-based additive,and a hydroxide source of an inorganic-based additive or anorganic-based additive to form precipitates comprising calciumsulfoaluminate or calcium sulfoferrate in a second precipitator unit;and (ii) removing precipitates from the de-brucited saline stream by anappropriate filtration unit to produce the de-scaled saline stream anddewatered precipitates.

This invention is not restricted to use in connection with oneparticular application. This invention can generally be used forde-scaling to produce useful and salable salts, and if required,de-salting saline streams. Further objects, novel features, andadvantages of the present invention will be apparent to those skilled inthe art upon examining the accompanying drawings and upon reading thefollowing description of the preferred embodiments, or may be learned bypractice of the invention.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 illustrates a simplified flow diagram for a Recycle-BrineMulti-Stage Flash (RB-MSF) desalination train.

FIG. 2 illustrates the solubility limit of calcium sulfate anhydrous asa function of temperature and sodium chloride concentration.

FIG. 3 illustrates a simplified flow diagram for a Once-ThroughMulti-Stage Flash (OT-MSF) desalination train.

FIG. 4 illustrates a simplified flow diagram for integrating thede-scaling step with a RB-MSF train.

FIG. 5 illustrates a simplified flow diagram for integrating thede-scaling step with the new concept of a Brine-Forward Multi-StageFlash (BF-MSF) desalination plant.

FIG. 6 illustrates a simplified flow diagram for integrating thede-scaling step with the new concept of a BF-MSF desalination plant withdirect-contact heat exchanging.

DESCRIPTION OF THE PREFERRED EMBODIMENT The Precipitation Concept

I have previously invented the Liquid-Phase Precipitation process (LPP)for the separation of inorganic species from saline streams. The effectof the separation in LPP is to intermix a saline stream with a suitablesolvent at ambient temperature and atmospheric pressure to formselective precipitates. The suitable solvents are those which have thecapability to meet two basic criteria.

The first criteria is the suitability to precipitate targeted inorganicspecies from saline streams. The selected organic solvent must bemiscible or at least soluble in water. Of equal importance, the targetedinorganic species must be sparingly soluble in the organic solvent. Theaddition of such a solvent to a saline stream leads to the capture ofpart of the water molecules and reduces the solubility of inorganicspecies in the water which form insoluble precipitates. The solubilityof the targeted inorganic species in the organic solvent is a criticalfactor in achieving the degree of saturation. Solubility related factorssuch as ionic charge, ionic radius, and the presence of a suitable anionin the saline stream play important roles in affecting andcharacterizing precipitates formation.

The second criteria is suitability for overall process design. For easeof recovery, the selected solvent must have favorable physicalproperties such as low boiling point, high vapor pressure, high relativevolatility, and no azeotrope formation with water. From a process designstandpoint, the selected solvent must have relatively low toxicity sincetraces of the organic solvent always remain in the discharge stream.Further, the selected solvent must be chemically stable, compatible withthe process, and relatively inexpensive.

The energy required to thermally separate the solvent from the aqueousmixture after precipitating the targeted inorganic species generallydepends on the solvent's boiling point. For distillation purposes,solvent's boiling point determines the number of degrees to which themixture must be heated. However, the solvent's specific heat and heat ofvaporization are also critical. The specific heat of the solventdetermines the number of calories that must be introduced into themixture to raise it each degree whereas the heat of vaporizationdetermines the number of additional calories needed to vaporize thesolvent. As such, the smaller the solvent's specific heat and heat ofvaporization, the fewer calories required for its thermal recovery.

Low boiling point solvents are thus preferred because the lessdifference between the mixture temperature and the solvent's boilingtemperature, the fewer calories required for thermally separating thesolvent from the mixture. In addition, with higher boiling pointsolvents, less complete solvents thermal recovery since the recoveredsolvents carry significant amounts of water. Carried over water, withits high heat of vaporization, represents an energy penalty.

Several amine solvents have been identified for potential use in the LPPprocess. The selected solvents, which are derivatives of ammonia, areprimary and secondary amines. They include methylamine (MA), ethylamine(EA), isopropylamine (IPA), propylamine (PA), dimethylamine (DMA),diethylamine (DEA), diisopropylamine (DIPA), and dipropylamine (DPA).

Nitrogen can form compounds with only three covalent bonds to otheratoms. An ammonia molecule contains sp³-hybridized nitrogen atom bondedto three hydrogen atoms. On the other hand, an amine molecule containssp³-hybridized nitrogen atom bonded to one carbon atom (primary amines)or more carbon atoms (2 carbon atoms in the case of secondary amines).The nitrogen has one orbital filled with a pair of unshared valenceelectrons, which allows these solvents to act as bases. Ammonia and theselected amines are therefore weak bases that could undergo reversiblereactions with water or other weak acids. However, when such solventsreact with a strong acid, their unshared electrons are used to formsigma bond with the acid, which drives the reaction to completion(irreversibly).

Table 3 presents some of the important characteristics of the selectedsolvents. However, IPA is the preferred solvent in the LPP process. Thepreference of using IPA is attributed to its high precipitation abilitywith different basic salts, overall favorable properties, near completereaction with strong acids, and relatively minimal environmental risks.

Improving the performance of LPP is always a target. One of theessential improvements is to minimize, if not eliminate, the use of theamine solvent. Inorganic additives can alternatively replace organicadditives or can be used in addition to organic additives to induceprecipitation of targeted species. The suitable inorganic additives forLPP are those that can form an insoluble inorganic-based mineral oftargeted charged species in a saline stream. Such inorganic additivesshould preferably be recoverable and recyclable, useable as a usefulby-product, or produced locally from reject or waste streams. Suchinorganic additives should also not themselves constitute pollutants.Several inorganic additives were indentified, developed, and tested forLPP.

A second targeted improvement for LPP is to produce controllableprecipitates that are uniformly distributed with high yield andpreferably in submicron sizes. Submicron precipitates are fundamentallystable and form spontaneously if a narrow resistance time distributionis improvised and/or a surface active agent (naturally existing orinduced) sufficiently acts as a dispersant to prevent immediateagglomeration of the newly formed precipitates. Submicron precipitatesare thus dispersed phase with extreme fluxionality. On the other hand,non-spontaneous unstable macro-size precipitates will form if givensufficient time to rest.

The state (stabile, metastabe, or unstable) of given precipitates can beexpressed thermodynamically by the Gibbs-Helmholtz relation as follows:ΔG=ΔH−TΔS  (1)where ΔG is the free energy of precipitates (provided by, for instance,mechanical agitation or other means), ΔH is the enthalpy that representsthe binding energy of the dispersed phase precipitates in the salinestream, T is the temperature, and ΔS is the entropy of the dispersedphase precipitates (the state of precipitates disorder). The bindingenergy (ΔH) can be expressed in terms of the surface tension (τ) and theincrease in the surface area (ΔA) as follows:ΔG=τΔA−TΔS  (2)When the introduced free energy into the saline stream exceeds thebinding energy of precipitates, individual precipitates are broken downand redistributed. In addition, when a surface active agent is presentin the saline stream as an effective dispersant, τ is reduced and thusthe precipitates binding energy is diminished. Furthermore, part of theintroduced energy may not contribute to precipitates deflocculating butit dissipates in the aqueous stream in the form of heat which reducesviscosity. All of these factors increase precipitates disorder (positiveentropy). As such, the change in the entropy (ΔS) quantitatively definesprecipitates dispersion.

The Compressed-Phase Precipitation (CPP) process is thus developed toachieve sub-micron precipitates in certain applications. CPP isconceptually similar to LPP in which the targeted inorganic species mustbe nearly insoluble in the amine solvent whereas the mother solvent(water) is miscible with the amine solvent. However, the difference isthat fluids in the CPP process can be subjected to pressure and/ortemperature manipulations, or fluids modifications to force unusualthermo-physical properties (e.g., exhibit liquid-like density but withhigher diffusivity, higher compressibility and lower viscosity).

The fast diffusion combined with low viscosity of a compressed aminesolvent into an aqueous phase produces faster supersaturation oftargeted ionic species, and their possible precipitation in the desiredand sub-micron and micron sizes. Thus, the precipitate-size as well asthe precipitate-size distribution, morphology, and crystal structure canbe controlled. Achieving faster supersaturation would, in turn, minimizethe use of the amine solvent, reduce the size of precipitation vessels,and allow the recovery of targeted ionic species in the desiredprecipitates shape and distribution.

Several factors could influence the performance of the precipitationprocess. Among such factors are: (1) the origin of the saline streamalong with the identity and concentrations of its targeted ionicspecies; and (2) the way the additive (inorganic, organic, or both) ispremixed or mixed with the saline stream to induce precipitation.

Treatment of Seawater and the Like of Saline Streams

De-Scaling Seawater and the Like of Saline Streams

Magnesium can be precipitated in the form of hydroxide (brucite) whereassulfate can be precipitated in the form of either calcium sulfoaluminateor calcium sulfoferrate. The precipitation of brucite and either calciumsulfoaluminate or calcium sulfoferrate can be conducted simultaneouslyin a single stage upon the addition of a hydroxide source as well as analuminum or iron source. The hydroxide source can be an inorganic-basedadditive, an organic-based additive, or a combination of such additives.

The precipitation of brucite and either calcium sulfoaluminate orcalcium sulfoferrate can preferably be conducted sequentially to recoverthem separately as salable and/or usable commodities. Brucite isprecipitated first upon the addition of an appropriate amount of ahydroxide source followed by the precipitation of either calciumsulfoaluminate or calcium sulfoferrate upon the addition of anappropriate amount of either an aluminum or iron source along with, ifneeded, an excess amount of a hydroxide source.

The possible hydroxide source of the inorganic-based additive is calciumchloroaluminate, calcium chloroferrate, lime, and hydrated lime. Thepossible hydroxide source of the organic-based additive is ammonia,methylamine, ethylamine, isopropylamine, propylamine, dimethylamine,diethylamine, diisopropylamine, and dipropylamine. The possible aluminumsource includes calcium chloroaluminate, aluminum chlorohydrate, calciumaluminate, sodium aluminate, aluminum acetate, aluminum chloride, andaluminum nitrate. The possible iron source includes calciumchloroferrate, calcium ferrate, sodium ferrate, iron chloride, and ironnitrate.

Calcium chloroaluminate and calcium chloroferrate are of particularinterest in this invention since they are layered double hydroxideswhich contain divalent and trivalent cations (calcium along withaluminum or iron) in the outside main layers while their interlayerscontain an anion (chloride) and water molecules. Within the outsidelayers, a fraction of calcium hydroxide sheets are substituted withaluminum or iron (trivalent cations), which provides permanent positivecharge on the hydroxide layers. The positively charged hydroxide layersare counter-balanced by the negatively charged chloride interlayers.

There are several advantages for using calcium chloroaluminate orcalcium chloroferrate in this invention. First, their interlayers arehighly exchangeable. Sulfate, as a divalent anion, would have a higheraffinity to replace chloride (as a monovalent anion) to bind withcalcium-aluminum or calcium-iron and thus form calcium sulfoaluminate orcalcium sulfoferrate. Second, they provide a dual source for bothdivalent (calcium) and trivalent (aluminum or iron) cations. Third, theyalso furnish the needed hydroxide ion for pH elevation. Fourth, they canbe produced from waste streams such as oil-gas fields produced water andthe like as given in some of my recent patent applications (Ser. Nos.13/066,841 and 13/507,141).

Integrating De-Scaling with De-Salting Seawater and the Like of SalineStreams

The de-scaling step is capable of depleting magnesium to about 1 mg/Land sulfate to about 40 mg/L (or even lower) from seawater. If thede-scaling step is integrated with seawater desalination plants, theremaining salts in seawater would be nearly in the form ofsodium-potassium chloride (sylvinite). With such an integration, anythermal-driven distillation method can achieve a reject brineconcentration (C_(B)) of 250,000 mg/L of TDS. Thermal-drivendistillation methods include, for instance, MSF, Multi-EffectDistillation (MED), Thermal Vapor Compression (TVC), Mechanical VaporCompression (MVC), Membrane Distillation (MD), or a combination of suchmethods. Pressure-driven (induced or osmotic) desalination methods suchas Reverse Osmosis (RO), Forward Osmosis (FO), and Osmotic MembraneDistillation (OMD) can also be efficiently used with the de-scaling stepto desalt seawater. However, RO must be operated at a much lower levelof C_(B) due to the osmotic pressure restriction. Design flow sheetsthat can be envisioned for conducting and optimizing each of suchdistillation methods or a combination of such methods to de-salt aneffectively de-scaled seawater and the like of saline streams are almostlimitless.

MSF is used as an example to demonstrate the benefits of the de-scalingstep of this invention. A modified RB-MSF train with the de-scaling stepis shown in FIG. 4. Pretreated (screening and chlorination) andpreheated seawater by the heat rejection stage of the RB-MSF train [10]is fed to the first precipitation unit [12] where it is intermixed witha hydroxide source [14] to form selective brucite precipitates. Thehydroxide source can be an inorganic-based additive, an organic-basedadditive, or a combination of such additives to form selective bruciteprecipitates. The outlet stream [16] from the first precipitation unit[12] is directed to the first filtration unit [18] to separate theformed brucite precipitates (slurry) [20] from the de-brucited seawater[22]. The brucite slurry [20] is subjected to further dewatering anddehydration (not shown).

The de-brucited seawater [22] is then fed to the second precipitationunit [24] where it is intermixed with an aluminum source or an ironsource [26] along with, if needed, an excess hydroxide source [14] toprecipitate either calcium sulfoaluminate (upon the addition of analuminum source) or calcium sulfoferrate (upon the addition of an ironsource). The outlet stream [28] from the second precipitation unit [24]is then fed to the second filtration unit [30] to separate the formedprecipitates (slurry) [32] from the de-sulfated seawater [34]. Theslurry of either calcium sulfoaluminate or calcium sulfoferrate [32] issubjected to further proper dewatering and dehydration (not shown).

A plurality of precipitator units as well as intermixing devices (e.g.,static mixers, premixing nozzles, concentric nozzles, spray nozzles,twin-fluid nozzles, Y-shaped nozzles, cross-shaped nozzles, or acombination of such nozzles) can be used in the de-brucitation andde-sulfation steps.

The de-scaled seawater [34] is then de-aerated by a stripping unit [36].The stripping unit can be a vacuum de-aerator or an indirect-contactheat exchanging stripper (a scrubber like unit). The indirect-contactheat exchanging stripper is of a particular interest since it not onlyfacilitates efficient removal of dissolved gases (gases' aqueoussolubilities generally decrease with increasing temperature) as well asefficient recovery the organic-based additive when used as a hydroxidesource, but also adds heating energy to the de-scaled seawater. Wasteheat is the preferred energy source that can be ideally paired with theindirect-contact heat exchanging stripper.

One of the sources of waste heat energy within power-seawaterdesalination co-generation plants is the returned condensates from brineheaters and low-pressure (LP) steam turbines. Cooling such pressurizedcondensates prior to their discharge into the atmospheric CondensatesRecovery Tank (CRT) is essential to prevent condensates from flashing.If condensates are allowed to flash (without cooling), their sensibleand latent heat energy will be lost in the CRT vent. Capturing the heatenergy from condensates and transferring it to the de-scaled seawatervia the indirect-contact heat exchanging stripper is ideal for heatingfurther the de-scaled seawater and stripping off dissolved gases andorganic-based additives.

When the hydroxide source is an inorganic-based additive, theindirect-contact heat exchanging stripper [36] as shown in FIG. 4 isused to strip off dissolved gases [38]. On the other hand, when thehydroxide source is an organic-based additive, the indirect-contact heatexchanging stripper [36] is used to strip off both the organic-basedadditive and dissolved gases. The stripped off organic-based additive isthen condensed (not shown) to segregate it [14] from the non-condensablegases [38] and recycled it to a storage tank (not shown) for reuse.

Seawater (pre-heated, de-scaled, de-aerated, and possibly furtherheated) is fed to the heat recovery section of the RB-MSF train. Thede-scaling step reduces the TDS in seawater from 49,950 mg/L (Table 1,2013 data) to about 40,000 mg/L. One of the possible ways to operate theRB-MSF plants are to fix C_(B) at 250,000 mg/L at the same flashingrange (70.1° C.) as each of the RB-MSF trains originally designed. C_(B)at 250,000 mg/L and 40° C. from the last stage of the RB-MSF train isstill well below the saturation limit of sodium chloride (about 365,000mg/L). As such, the unsaturated but highly concentrated sodium chlorideat 250,000 mg/L is unlikely to form precipitate films in the vicinity ofheat transfer (condensers/pre-heaters) tubing. The pressure in the firstflashing stage of the heat recovery section should be 768.8 mmHg (theboiling point of recycle brine is 106.93° C.) and decreases gradually tolower the boiling point of the flashed off brine in each consecutivestage until it reaches 42.8 mmHg in the last flashing stage of the heatrejection section (C_(B): 250,000 mg/L; and T_(B): 39.9° C.).

Table 4 presents the simulated performance of the de-scaled RB-MSF trainwhen the C_(B) and flashing range are fixed at, respectively, 250,000mg/L and 70.1° C. to maintain the same production of distillate (593kg/s or 15 MGD) as the original RB-MSF train. The recycle brine flowrate (Q_(r)) remains the same (5,398 kg/s or 136.6 MGD) as the originalRB-MSF train without the de-scaling step. The cooling seawater flow rate(Q_(c)) is increased to 4,053 kg/s (102.6 MGD) in the de-scaled RB-MSFfrom 2,720 kg/s (68.8 MGD) in the original RB-MSF. However, seawaterflow rate (Q_(F)) to be mixed with flashed off brine from the last stageof the heat rejection section to form recycle brine to feed the heatrecovery section is substantially reduced from 2,829 kg/s (71.6 MGD) inthe original RB-MSF train to 706 kg/s (17.9 MGD) in the de-scaled RB-MSFtrain. Reject brine flow rate (Q_(B)) is also substantially reduced from2,236 kg/s (57.6 MGD) in the original RB-MSF train to 113 kg/s (2.9 MGD)in the de-scaled RB-MSF train. Thus, the total reduction in Q_(F) forthe de-scaled RB-MSF Plants I and II (8 trains per plant) would be 859MGD whereas the total reduction in Q_(B) from both plants would be 875MGD. This represents tremendous savings in pumping power, de-scalingadditives, down sizing de-aeration equipment.

It should be pointed out that the total required seawater (Q_(Fc):retuned cooling seawater+seawater to be mixed with flashed off brine toform recycle brine) per train is reduced to 4,760 kg/s (120.5 MGD) inthe de-scaled RB-MSF train from 5,549 kg/s (140.4 MGD) in the originalRB-MSF train. The increase in returned cooling seawater in the de-scaledRB-MSF train is attributed to the reduction of seawater to be mixed withflashed off brine to form recycle brine. However, returned coolingseawater is not subjected to the de-scaling step and it is pre-treatmentis only limited to screening and chlorination as well as a portion of it(about 20%) is recycled for bending with seawater before entering theheat rejection section.

The designed top brine temperature (T_(TB)) range of MSF plants istypically 120-150° C. For such a T_(TB) range, steam to derive brineheaters of MSF trains can be extracted from the steam turbines stage(the heat recovery circuits of the coupled power plants) between thehigh-pressure (HP) and intermediate-pressure (LP) steam turbines or fromthe HP steam turbine. It provides higher temperature flashing ranges andallows expanding the heat recovery section of the MSF train from 20 to40 stages. This, in turn, allows better performance ratio (e.g., 20),which reduces the steam load and/or increases distillate production aswell as allows steady operation (due to better performance of heattransfer coefficients). Such a higher T_(TB) range is mainly avoidedbecause of the calcium sulfate scale envelop and the temperaturetolerance of scale inhibitors is limited to 110° C. The de-scaling stepwould thus lift such imposed restrictions on T_(TB).

Table 4 also shows the simulated performance of the de-scaled RB-MSFtrain when C_(B) and flashing range are fixed at, respectively, 250,000mg/L and 110.1° C. (T_(TB): 150° C.) to maintain the same production ofdistillate (593 kg/s or 15 MGD). Q_(F) and Q_(B) also remain the same asin the lower flashing range. However, the higher flashing range (110.1°C.) substantially reduces Q_(r), Q_(c), Q_(F), and Q_(S) compared to thelower flashing range (70.1° C.) in each train as follows: (1) Q_(r) isreduced to 2,886 kg/s (73.0 MGD) from 5,398 kg/s (136.6 MGD); (2) Q_(r)is reduced to 3,384 kg/s (85.6 MGD) from 4,053 kg/s (102.6 MGD); (3)Q_(Fc) is reduced to 4,090 kg/s (103.5 MGD) from 4,760 kg/s (120.5 MGD);and (4) Q_(S) is reduced to 59.6 kg/s from 83.3 kg/s. The reduction inQ_(r) reduces the recycle brine concentration (C_(r)) to 198,622 mg/Lfrom 222,513 mg/L whereas the reduction in Q_(S) increases theperformance ratio from 7.1 to 10. For the de-scaled RB-MSF Plants I andII (8 trains per plant) at the higher flashing range compared to thelower flashing range, the total reduction in: (1) Q_(r) is 40,192 kg/s(1,017.1 MGD); (2) Q_(c) is 10,700 kg/s (270.8 MGD); (3) Q_(Fc) is10,720 kg/s (271.3 MGD); and (4) Q_(S) is 379.2 kg/s. Hence, thede-scaling step allows RB-MSF plants to reach their designed T_(TB)(150° C.) and substantially reduces Q_(r), Q_(c), Q_(Fc) and Q_(S),which tremendously reduces pumping power and increases the performanceratio. As such, two RB-MSF trains with one heat rejection section can becoupled with one brine heater or PR can be doubled by using a train of43-stages rather than 23-stages.

The de-scaling step clearly aids RB-MSF plants in resolving calciumsulfate scale, increasing PR, and substantially reducing Q_(F), Q_(r),Q_(c), Q_(Fc) and Q_(B). However, the required pumping power forreturned cooling seawater and recycle brine still remains significant,which is the inherited disadvantage of the RB-MSF concept. Anotherstriking disadvantage of the RB-MSF concept, which stems from need forcooling seawater in the heat rejection section of each train, is thatthe de-scaling step must be added to each train of the RB-MSF plant. Toresolve this issue, separate cooling seawater can be used for the heatrejection section of each train, while seawater to be mixed with recyclebrine can be processed independently using one de-scaling step coupledwith one indirect-contact heat exchanging stripper to de-scale, preheatand de-aerate seawater before feeding the whole RB-MSF plant.

Alternatively, what I coined Brine-Forward MSF (BF-MSF) desalinationplant in conjunction with the de-scaling step as depicted in FIG. 5,would entirely eliminate recycle brine and require only one de-scalingstep for the whole plant. For a comparison purpose, the BF-MSF plantconsists of eight trains and each train consists of total 23 stages(similar to the recently installed RB-MSF plants in terms of number oftrains and total number of stages per train). Each train of the BF-MSFplant consists of 23 stages of heat recovery except the last train,which is divided into a heat recovery section (20 stages) and a heatrejection section (3 stages) to recover more distillate from the lasttrain at low temperatures (40-33° C.) before rejecting brine. Coolingseawater is only used for the heat rejection section of the last train.Cooling seawater (Q_(c)) that gains heat as it exists the heat rejectionsection of the last train [10A] is completely used (no rejected Q_(c))by mixing with pretreated seawater (Q_(F)) [10B] to partially preheatand compose the total required seawater (Q_(Fc)) [10]. Q_(Fc) in theBF-MSF plant is thus the sum of Q_(c) and Q_(F), which is totallyutilized to feed the trains. This is unlike the RB-MSF plant whereQ_(Fc) is the sum of the utilized Q_(c) in feeding the trains (referredto as Q_(F)) and the returned (rejected) Q_(c).

The partially preheated total required seawater in the BF-MSF plant asshown in FIG. 5 is subjected to de-scaling and de-aeration (similar tothe processing steps [10-38] of FIG. 4 as described above) beforeentering the first train. The brine forward feeding concept is thenapplied, which is the brine from each train passes through to feed thenext train of the BF-MSF plant, and each train possesses a higherconcentration factor (C_(B)/C_(F): 1.26) of brine than the previoustrain. This allows consecutive increase in the concentration factor ofbrine as it gets more concentrated in going down the trains. TDS inreject brine (C_(B)) exiting the last train can be set at 250,000 mg/L.

The BF-MSF concept is thus multi-boiling without supplying additionalheat after the train's brine heater, and multi-concentration withoutsupplying additional seawater after the first train of the plant. Theconcept is aimed at entirely eliminating recycle brine, substantiallyreducing as well as completely utilizing cooling seawater, and using onede-scaling step, rather than multiple de-scaling steps, to serve thewhole plant.

Table 5 presents the simulated performance of the de-scaled BF-MSF plant(FIG. 5) at 250,000 mg/L of C_(B) with a lower flashing range (70.1° C.)and a higher flashing range (110.1° C.). At the lower flashing range,the de-scaled BF-MSF plant provides higher PR (8.23) but lower Q_(D)(3,742 kg/s or 94.7 MGD) and higher Q_(B) (1,906 kg/s or 48.2 MGD) thanthe de-scaled RB-MSF plant (PR: 7.12; Q_(D): 4,744 kg/s or 120 MGD; andQ_(B): 904 kg/s or 22.9 MGD) at the same flow rate of seawater thatactually feeds the trains (Q_(FC) in the case of BF-MSF and Q_(F) in thecase of RB-MSF). However, the distinct advantages of the de-scaledBF-MSF plant compared to the de-scaled BR-MSF plant (Table 4) are: (1)the total elimination of Q_(r); and (2) the substantial reduction ofQ_(c) (364 kg/s or 9.2 MGD) as well as its complete utilization(returned Q_(c) from the BR-MSF plant is 32,422 kg/s or 820.5 MGD).

At the higher flashing range, the de-scaled BF-MSF plant providesequivalent performance to the de-scaled RB-MSF plant in terms of Q_(D),Q_(B), Q_(S) and PR. Furthermore, the de-scaled BF-MSF plant eliminatesQ_(r), significantly reduces Q_(c) to 473 kg/s (12.0 MGD), andcompletely utilizes Q_(c) compared to the de-scaled BR-MSF plant(returned Q_(c) from the BR-MSF plant is 27,072 kg/s or 685.1 MGD). Themain objective of the de-scaling step is therefore achieved by allowinga higher flashing range and a better MSF design (BF concept) that doesnot require pumping enormous volumes of recycle brine and coolingseawater.

Sufficient seawater preheating is essential, but is not a must, sinceonly the rear-end stages of the first train of the BF-MSF plant may beaffected by it. If the striping unit is not based on indirect-contactheat exchanging or returned condensates from brine heaters and/or LPsteam turbines are not available to further heat the de-scaled seawater,then one of the possible ways to sufficiently preheat seawater is toutilize blow down streams from brine heaters and/or from the LP steamturbines. Blow down is necessary to control dissolved solids andscale-prone species in brine heaters and steam turbines, yet blow downstreams from brine heaters and steam turbines account for 3% loss oftheir energy input. Significant amounts of heat input to the brineheaters and steam turbines are lost as both sensible and latent heatenergy contained in the existing blow down streams at 110-120° C. (atthe lower MSF flashing range). Heat recovery from such blow down streamsby a condensing heat exchanger can offset much, if not all, of therequired thermal energy to sufficiently preheat seawater. Adirect-contact heat exchanger is ideal for removing heat from such blowdown streams to preheat seawater.

FIG. 6 shows the integrated de-scaling step with the BF-MSF plant usingdirect-contact heat exchanging to preheat seawater. Cooling seawaterthat gains heat as it exists the heat rejection section of the lasttrain [10A] is mixed with pretreated and preheated seawater [10B] tocompose the preheated total required seawater [10]. The preheating ofseawater [10B] is accomplished by a direct-contact heat exchanger [10C]using blow down streams [10D] from brine heaters and/or LP steamturbines. The de-scaling step facilitates the efficient use of thedirect-contact heat exchanger since it's located after preheating takesplace and it targets the precipitation of scale-prone species from thepreheated total required seawater, which is a mixture of seawater, blowdown streams and cooling seawater. The de-scaling and de-aeration of thepreheated total required seawater as shown in FIG. 6 are similar to theabove described processing steps [10-38] in FIG. 4. The preheated,de-scaled and de-aerated seawater is then fed to the first train of theBF-MSF plant where the brine forward feeding concept is applied tosubsequent trains.

The de-scaling step also facilitates the pairing of two BF-MSF heatrecovery trains with one brine heater or expanding the stages of eachheat recovery train to 40 stages at the higher flashing range (100.1°C.).

The integrated RB-MSF or BF-MSF plant with the de-scaling step canreject brine at lower C_(B) levels. As given in one my newly filedpatent applications, the preference of rejecting brine at 250,000 mg/Lis that rather than blowing it back to the sea, it can be utilized toeffectively precipitate and produce sodium carbonate in profitableamounts at such C_(B) in conjunction with recovering CO₂ and waste heatenergy from flue gas of the coupled power plants. The recovered wasteheat energy from flue gas can offset the steam required at high flashingranges. This offset will reduce fuel consumption while maintaining afixed net steam output from the coupled “combined cycle” power plant, orwhen required, it will increase the net steam output by maintaining thesame fixed fuel consumption.

Reject brine from RB-MSF or BR-MSF plants can also be used foroil-fields Improved Oil Recovery (IOR) as an ideal fluid since itsreadily depleted of sulfate, other scale-prone species, and oxygen.

TABLE 1 Changes in Seawater before and after Operating BR-MSF Plants.2006 2009 2013 Ions Before Operating After Operating After Operating(mg/L) Plant I Plant I Plants I & II Na⁺ 12,173 13,630 14,530 K⁺ 423 524611 Mg⁺² 1,529 1,886 2,145 Ca⁺² 530 667 699 Sr⁺² 5 7 8 Cl ⁻ 22,00025,550 27,710 HCO₃ ⁻ 140 146 147 SO₄ ⁻² 2,700 3,770 4,100 TDS 39,50046,200 49,950

TABLE 2 BR-MSF vs. OT-MSF Performance Based on Seawater Variations.Start-up Fixed C.F. Fixed C_(B) Fixed C_(B) RB-MSF RB-MSF RB-MSF OT-MSFParameters (2006) (2013) (2013) (2013) MSF Train: C_(F): TDS, mg/L39,500 49,950 49,950 49,950 C_(F): SO₄ ⁻², mg/L 2,700 4,100 4,100 4,100T_(F), ° C. 33 33 33 33 T_(B), ° C. 39.9 39.9 39.9 39.9 T_(TB), ° C. 110110 110 110 #S 23 23 23 23 Q_(F), kg/s 1,617 1,617 2,829 2,829 Q_(c),kg/s 3,932 3,932 2,720 0 Q_(Fc), kg/s 5,549 5,549 5,549 2,829 Q_(r),kg/s 5,398 5,398 5,398 0 C_(r): TDS, mg/L 56,101 70,943 56,255 0 C_(r):SO⁻² ₄, mg/L 3,835 5,823 4,618 0 C_(r)/C_(F) 1.42 1.42 1.13 0 Q_(B),kg/s 1,011 1,011 2,236 2,236 C_(B): TDS, mg/L 63,200 79,920 63,20063,200 C_(B): SO⁻² ₄, mg/L 4,320 6,560 5,188 5,188 C_(B)/C_(F) 1.6 1.61.27 1.27 Q_(D), kg/s 593 593 593 593 Q_(S), kg/s 62.9 62.9 62.9 33 PR9.4 9.4 9.4 13.6 MSF Plant (8 Trains): ΣQ_(F), kg/s (MGD) 12,936 (327.4)12,936 (327.4)  22,632 (572.8)  22,632 (572.8) ΣQ_(c), kg/s (MGD) 31,456(769.1) 31,456 (769.1)  21,760 (550.7)   0 ΣQ_(Fc), kg/s (MGD)  44,392(1,123.4) 44,392 (1,123.4) 44,392 (1,123.4) 22,632 (572.8) ΣQ_(r), kg/s(MGD)  43,184 (1,092.9) 43,184 (1,092.9) 43,184 (1,092.9) 0 ΣQ_(B), kg/s(MGD)  8,088 (204.7) 8,088 (204.7) 17,888 (452.7)  17,888 (452.7)ΣQ_(D), kg/s (MGD) 4,744 (120) 4,744 (120)  4,744 (120)  4,744 (120)ΣQ_(S), kg/s (MGD) 503.2 503.2 503.2 264.0 C.F.: Concentration Factor;Q_(F): Seawater Feed Mass Flow to be mixed with flashed off brine;Q_(c): Cooling Seawater Mass Flow; Q_(Fc): Total Required Seawater MassFlow (Q_(F) + Q_(c)); C_(F): Seawater Concentration; T_(F): SeawaterTemperature; T_(B): Reject Brine Temperature; T_(TB): Top BrineTemperature; #S: Number of Stages; Q_(r): Recycle Brine Mass Flow; C_(r): Recycle Brine Concentration; Q_(B): Reject Brine Mass Flow; C_(B):Reject Brine Concentration; Q_(D): Distillate Mass Flow; Q_(S) : SteamMass Flow; PR: Performance Ratio; and MGD: million US gallons per day.

TABLE 3 Properties of Selected Pure Solvents. Solubility T_(b) C_(p)H_(v) ΔH_(f) ⁰ Fluid in Water ° C. kJ/kg ° C. kJ/kg kJ/kg NH₃ HighlySoluble −33.45 2.19 1,370.8 −2,695 MA ( CH₅N ) Extremely Soluble −6.353.28 790.8 −1,517 DMA(C₂H₇N) Extremely Soluble 6.85 3.03 587.4 −974 EA(C₄H₁₁N) Miscible 16.55 2.85 621.8 −1,644 DEA (C₄H₁₁N) Miscible 55.452.44 380.4 −1,418 IPA (C₃H₉N) Miscible 32.45 2.77 460.1 −1,900 DIPA(C₆H₁₅N) Highly Soluble 83.95 2.64 341.9 −1,765 PA (C₃H₉N) Miscible48.65 2.75 502.6 −1,717 DPA (C₆H₁₅N) Soluble 109.35 368.6 −1,543 T_(b):Normal Boiling Point; C_(p): Specific Heat Capacity; H_(v): Heat ofVaporization at Normal Boiling Point; and ΔH_(f) ⁰: Standard Enthalpy ofFormation.

TABLE 4 Integrating the De-Scaling Step with the BR-MSF Plant. LowerFlashing Higher Parameters Range as designed Flashing Range MSF Train:C_(F): TDS, mg/L 40,000 40,000 C_(F): SO₄ ⁻², mg/L 40 40 T_(F), ° C. 3333 T_(B), ° C. 39.9 39.9 T_(TB), ° C. 110 150 #S 23 23 Q_(F), kg/s 706706 Q_(c), kg/s 4,053 3,384 Q_(Fc), kg/s 4,760 4,090 Q_(r), kg/s 5,3982,886 C_(r): TDS, mg/L 222,513 198,622 C_(r): SO₄ ⁻², mg/L 222.5 198.6C_(r)/C_(F) 5.56 4.97 Q_(B), kg/s 113 113 C_(B): TDS, mg/L 250,000250,000 C_(B): SO₄ ⁻², mg/L 250 250 C_(B)/C_(F) 6.25 6.25 Q_(D), kg/s593 593 Q_(S), kg/s 83.3 59.6 PR 7.12 9.95 MSF Plant (8 Trains): ΣQ_(F),kg/s (MGD) 5,648 (142.9) 5,648 (142.9) ΣQ_(c), kg/s (MGD) 32,422 (820.5)27,072 (685.1) ΣQ_(Fc), kg/s (MGD) 38,080 (963.7) 32,720 (828.1) ΣQ_(r),kg/s (MGD)  43,184 (1,092.9) 23,088 (584.3) ΣQ_(B), kg/s (MGD)  904(22.9)   904 (22.9) ΣQ_(D), kg/s (MGD) 4,744 (120)  4,744 (120) ΣQ_(S),kg/s (MGD) 666.4 476.8 PR 7.12 9.95

TABLE 5 Integrating the De-Scaling Step with the Design BF-MSF Plant.Lower Flashing Higher Flashing Parameters Range Range C_(F): TDS, mg/L40,000 40,000 C_(F): SO₄ ⁻², mg/L 40 40 T_(F), ° C. 33 33 T_(B), ° C.39.9 39.9 T_(TB), ° C. 110 150 #S 23 23 BF-MSF Train #1: Q_(F), kg/s5648 5648 Q_(D1), kg/s 717.2 1159.4 Q_(S1), kg/s 87.2 116.6 Q_(B1), kg/s4,931 4,488.2 C_(B1)/C_(F) 1.26 1.26 BF-MSF Train #2: Q_(D2), kg/s 626.1921.4 Q_(S2), kg/s 76.1 92.7 Q_(B2), kg/s 4,305 3566.8 C_(B2)/C_(B1)1.26 1.26 BF-MSF Train #3: Q_(D3), kg/s 546.6 732.3 Q_(S3), kg/s 66.473.6 Q_(B3), kg/s 3,758 2,834.5 C_(B3)/C_(B2) 1.26 1.26 BF-MSF Train #4:Q_(D4), kg/s 477.2 591.9 Q_(S4), kg/s 58.0 58.5 Q_(B4), kg/s 3,2812,252.6 C_(B4)/C_(B3) 1.26 1.26 BF-MSF Train #5: Q_(D5), kg/s 416.6462.5 Q_(S5), kg/s 50.6 46.5 Q_(B5), kg/s 2,864 1,790.1 C_(B5)/C_(B4)1.26 1.26 BF-MSF Train #6: Q_(D6), kg/s 363.7 367.5 Q_(S6), kg/s 44.237.0 Q_(B6), kg/s 2,501 1,422.6 C_(B6)/C_(B5) 1.26 1.26 BF-MSF Train #7:Q_(D7), kg/s 317.5 292.1 Q_(S7), kg/s 38.6 29.4 Q_(B7), kg/s 2,1831,130.5 C_(B7)/C_(B6) 1.26 1.26 BF-MSF Train #8: Q_(D8), kg/s 277.2232.1 Q_(S8), kg/s 33.7 23.3 Q_(B8), kg/s 1,906 898.5 Qc₈, kg/s 364 472C_(B8)/C_(B7) 1.24 1.24 BF-MSF Plant (8 Trains): Q_(F), kg/s 5,284(133.7) 5,175 (130.9) Q_(c), kg/s 364 (9.2) 473 (12.0) Q_(Fc), kg/s5,648 (142.9) 5,648 (142.9) ΣQ_(D), kg/s (MGD) 3,742 (94.7)  4,750(120.2) ΣQ_(S), kg/s (MGD) 454.7 477.6 PR 8.23 9.95 Q_(B), kg/s 1,906(48.2)  898 (22.7) C_(B): TDS, mg/L 250,078 250,078 C_(B): SO₄ ⁻², mg/L250 250 C_(B)/C_(F) 6.25 6.25

What is claimed is:
 1. A method for desalinating a de-scaled andde-aerated saline stream, said method comprising producing distillateand de-scaled brine in a desalination system, said desalination systemcomprises a Brine-Forward Multi-Stage Flash (BF-MSF) desalinationsystem, said BF-MSF desalination system comprises a plurality of trainsarranges in series, in which each train comprises a brine heater and aheat recovery section except a last train, said last train comprises abrine heater, a heat recovery section and a heat rejection section,wherein a de-scaled and de-aerated saline stream is fed to a first trainof said BF-MSF desalination system and wherein each train is operated toproduce a pre-selected concentration of distillate and de-scaled brine;wherein said de-scaled brine from said each train except said last trainpasses as a feed stream to the next succeeding train thereby producingan outlet stream having a progressively higher level of total dissolvedsolids (TDS) than a preceding train; wherein said de-scaled brine fromsaid last train is rejected having a TDS concentration not exceeding250,000 mg/L of said TDS; wherein said BF-MSF desalination system is amulti-boiling system which does not require additional heat afterpassing through said brine heater of each train; and amulti-concentration system which does not require supplying additionalde-scaled and de-aerated saline stream after said first train.
 2. Themethod of claim 1, wherein said de-scaled and de-aerated saline streamcomprising magnesium and sulfate is separated by (a) mixing said salinefeed stream with a hydroxide source to form a first precipitatecomprising brucite in a first precipitator unit; filtering said firstprecipitate by a first filter to produce a de-brucited saline stream;(b) separating said sulfate from said de-brucited saline stream bymixing said de-brucited saline stream with a trivalent cation source toform a second precipitate comprising either calcium sulfoaluminate orcalcium sulfoferrate in a second precipitator unit; followed byfiltering said second precipitate by a second filter to produce ade-scaled saline stream; and (c) de-aerating said de-scaled salinestream of step (b) using a stripping unit to produce a de-scaled andde-aerated saline stream free of magnesium and sulfate.
 3. The method ofclaim 2, wherein said saline feed stream is selected from the groupconsisting of seawater, brine from seawater desalination plants, naturalbrine, brackish water, produced water, flue gas desulphurization spentwater, agricultural drainage water, acid mine drainage water, mineralslurry transport water, paper mills spent water, aluminum anodizingspent water, spent water from fertilizer production, lime slaking, spentwater from textile production, and combinations thereof.
 4. The methodof claim 2 wherein said hydroxide source is selected from the groupconsisting of calcium chloroaluminate, calcium chloroferrate, lime,hydrated lime, ammonia, methylamine, ethylamine, isopropylamine,propylamine, dimethylamine, diethylamine, diisopropylamine,dipropylamine, and combinations thereof.
 5. The method of claim 2,wherein said trivalent cation source comprises either an aluminum-basedadditive or an iron-based additive.
 6. The method of claim 5, whereinsaid aluminum-based additive is selected from the group consisting ofcalcium chloroaluminate, aluminum chlorohydrate, calcium aluminate,sodium aluminate, aluminum acetate, aluminum chloride, aluminum nitrate,and combinations thereof.
 7. The method of claim 5, wherein saidiron-based additive is selected from the group consisting of calciumchloroferrate, calcium ferrate, sodium ferrate, iron chloride, ironnitrate, and combinations thereof.
 8. The method 2, wherein in step (b)said de-brucited saline stream is further mixed with a hydroxide source.9. The method of claim 2, wherein said stripping unit further comprisesan indirect contact heat exchanger to heat said de-scaled saline streamby a heat source.
 10. The method of claim 9, wherein said heat source isselected from the group consisting of a return condensate from a brineheater, a return condensate from a turbine and combinations thereof. 11.The method of claim 1, wherein said feed stream to the MSF DistillationSystem for producing a de-scaled and de-aerated saline stream ispre-heated.
 12. The method of claim 11, wherein the pre-heating of thefeed stream comprises the steps of passing at least a portion of saidsaline feed stream through heat transfer tubings of a heat rejectionsection in a train of a desalination system to exchange heat withflashing brine in said heat rejection section to produce a preheatedportion of saline feed stream.
 13. The method of claim 11, whereinpre-heating of a portion of said feed stream further comprise blendingwith at least a second portion of saline feed stream.
 14. The method ofclaim 13, wherein the second portion of saline feed stream is heatexchanged with a heat source to produce a second preheated saline feedstream.
 15. The method of claim 14, wherein said heat source is selectedfrom the group consisting of a blowdown stream from a brine heater, ablowdown stream from a turbine and combinations thereof.
 16. The methodof claim 14, wherein said feed stream is heat exchanged with a directcontact heat exchanger.
 17. The method of claim 1, wherein said BF-MSFdesalination system further comprises a plurality of trains arranged inseries, in which each train comprises a brine heater and a heat recoverysection, said de-scaled and de-aerated saline stream is fed to a firsttrain of said BF-MSF desalination system, said each train is operated ata preselected concentration and wherein each of train produces saiddistillated and said de-scaled brine stream, said de-scaled brine streamof each of said train except the last train passes through as the feedto next succeeding train whereby each train progressively produces anoutlet stream having a higher level of total dissolved solids (TDS) thanthe preceding train; and wherein said de-scaled brine from said lasttrain is rejected having a TDS concentration not to exceed 250.00 mg/Lof said TDS; wherein said BF-MSF desalination system is a multi-boilingsystem not requiring additional heat after passing through said brineheater of each train, and a multi-concentration system without supplyingadditional de-scaled and de-aerated saline stream after said firsttrain.
 18. The method of claim 1, wherein said desalination system isfurther selected from the group consisting of a recycle-brinemulti-stage flash desalination train, a once-through multi-stage flashdesalination train, multi-effect distillation, thermal vaporcompression, mechanical vapor compression, vacuum membrane distillation,direct-contact membrane distillation, osmotic membrane distillation,reverse osmosis, forward osmosis, and combinations thereof.